WO2014041285A1

WO2014041285A1 - Method for producing kerosene from butanols

Info

Publication number: WO2014041285A1
Application number: PCT/FR2013/052054
Authority: WO
Inventors: Tom VIVIEN; Vincent Coupard; Francois Hugues; Raphael Huyghe
Original assignee: IFP Energies Nouvelles
Priority date: 2012-09-12
Filing date: 2013-09-06
Publication date: 2014-03-20
Also published as: BR112015005125A2; EP2895448B1; FR2995306A1; US20150246855A1; US9670112B2; FR2995306B1; BR112015005125B1; EP2895448A1

Abstract

The invention concerns a method for producing middle distillate hydrocarbon bases from a butanol feedstock, said method comprising at least: a) a step of isomerisation dehydration of said butanol feedstock, b) a step of separating the water present in said butylene effluent, c) a step of purifying the organic liquid effluent from step b), d) a first step of oligomerisation of a feedstock comprising at least a portion of the purified organic effluent from step c), all of the effluent from step g) and at least a portion of the light product from step f), e) a second step of oligomerisation of said first oligomerisation effluent so as to produce a second oligomerisation effluent, f) a step of fractionation of said second oligomerisation effluent into at least three products corresponding respectively to a light product mostly comprising compounds C₂ to C4, an intermediate product mostly comprising compounds C5 to C9, and a middle distillate product mostly comprising the compounds having at least 10 carbon atoms, g) a third step of oligomerisation of at least a portion of said intermediate product, h) a step of hydrogenating at least a portion of said middle distillate product.

Description

PROCESS FOR PRODUCING KEROSENE FROM BUTANOLS

TECHNICAL FIELD OF THE INVENTION

The present invention relates to the conversion of butanol, and more particularly of biobutanol, to a fuel base.

It relates more particularly to a flexible catalytic process for converting butanol into middle distillates.

PRIOR ART

Demand for the use of renewable resources as a partial replacement of petroleum resources for fuel synthesis is increasing. Thus the use of biobutanol in the synthesis of bases for fuels is receiving more and more interest.

By biobutanol is meant a butanol produced from renewable resources derived from biomass such as, for example, lignocellulosic raw materials.

The latter are cellulosic materials, that is to say constituted more than 90% by weight of cellulose, and / or lignocellulosic, that is to say consisting of cellulose, hemicellulose, which are essentially constituted polysaccharides pentoses and hexoses as well as lignin, which is a macromolecule of complex structure and high molecular weight, composed of aromatic alcohols connected by ether bonds.

The emergence of biobutanol production projects is largely driven by the public and industrial will to develop second-generation biofuels, ultimately valuing a wide range of lignocellulosic raw materials.

In comparison with ethanol, n-butanol and isobutanol have several advantages: higher energy density; less hydrophilic character; better compatibility with both fuel storage / transport infrastructure and common automotive engines; lower vapor pressures; less corrosive. Isobutanol is particularly compatible with diesel fuel while n-butanol is miscible with gasoline.

N-Butanol is one of the products of so-called ABE (acetone-butanol-ethanol) fermentations of Clostridium acetobutylicum. Recent projects use selected strains, mutants or GMOs allowing a specific production of n-butanol, one of the limiting factors being its toxicity for the microorganisms concerned. More tolerant strains and continuous or semi-continuous extraction processes are therefore necessary. The bacteria used in these processes are always of the Clostridium genus. Isobutanol, or 2-methylpropane-1-ol, is a branched isomer of n-butanol. It is also a product of the fermentation of carbohydrates. Its biosynthetic pathways are however different from those of n-butanol. The processes for the production of isobutanol are being developed by actors such as Gevo or Butamax from bacterial sources different from those used for the production of n-butanol (E.Coli).

The patent application WO2009 / 079213 describes the production of bases for fuels from biomass comprising alcohol fermentation stages, dehydration of a portion of said alcohols to olefins, oligomerization of a part of said olefins and hydrogenation eventual product of the oligomerization. In contrast, the patent does not provide a means of controlling the temperature increase in the reactor due to the exothermicity of the oligomerization reaction. Oligomerization is carried out in a single step.

The patent application WO201 1/140560 describes the production of a kerosene base from lignocellulosic raw materials through an isobutanol route. Said raw materials are fermented under specific conditions specific to the production of isobutanol. This is then dehydrated and then oligomerized to produce a close cut of kerosene. This application does not address the problem of controlling the increase in temperature in the reactor due to the exothermicity of the oligomerization, nor the control of the composition of the products obtained.

OBJECT AND INTEREST OF THE INVENTION

The present invention relates to a process for producing medium distillate hydrocarbon bases from a butanol feedstock, said process comprising at least: a) a step of isomerizing dehydration of said butanol feedstock in the presence of an amorphous or zeolite acid catalyst in at least one reactor, operating at an absolute pressure at the reactor inlet of between 0.5 and 1.2 MPa and at a reactor inlet temperature of between 350 and 450 ° C. so as to produce a predominantly butylenic effluent, b ) a step of separating the water present in said butylenic effluent operating at a pressure of between 0.5 and 1.2 MPa and at a temperature of between 35 and 60 ° C., c) a step of purifying the effluent organic liquid from step b) to produce a purified organic effluent, d) a first oligomerization step of a feed comprising at least a portion of the organic effluent e purified from step c), all of the effluent from step g) and at least a portion of the light product from step f) in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 0.5 and 10 MPa, at a temperature of between 40 and 110 ° C. and at an hourly space velocity of between 0.1 and 10 h ^-1 , so as to produce an effluent of first oligomerization, comprising at least 50% by weight of olefins having a number of carbon atoms greater than or equal to 8, the weight percentage being expressed relative to the total mass of the olefins contained in said effluent, e) a second stage oligomerising said effluent of first oligomerization, in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 2 and 15 MPa, at a temperature of between 100 and 200 ° C. and at an hourly volume rate of between between 0.1 and 10 h ^-1 , so as to produce a second oligomerization effluent, f) a fractionation step of said second oligomerization effluent in at least three corresponding products respecti to a light product comprising predominantly C ₂ to C ₄ compounds, an intermediate product comprising predominantly C ₅ to C ₉ compounds, and a middle distillate product comprising predominantly compounds having at least 10 carbon atoms, g) a third step oligomerising at least a portion of said intermediate product in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 2 and 15 MPa, at a temperature between 100 and 200 ° C and at a rate of hourly volume of between 0.1 and 5 h ^-1 , h) a step of hydrogenation of at least a portion of said middle distillate product in the presence of a catalyst comprising at least one Group VI II metal in at least one reactor operating at an absolute pressure of between 2 and 4 MPa, at a temperature of between 100 and 350 ° C. and at an hourly volume velocity of between 1 and 5 h ^-1 with a molar ratio of hydrogen s ur hydrocarbons H ₂ / HC between 10 and 450.

An advantage of the invention is a better control of the products obtained, by the separation of the two oligomerization steps, said separation also making it possible to limit the increase in temperature in the reactor due to the exothermicity of the reactions in each of said steps.

Another advantage of the invention is that aromatics can be added to the fuel base produced in accordance with the invention up to the limits provided by ASTM D7566-1 1a while respecting the density specification of said standard. .

DETAILED DESCRIPTION OF THE INVENTION Charge

The filler treated in the process according to the invention is a butanol filler, the term butanol characterizing all the alcohols comprising 4 carbon atoms. Said butanol filler is advantageously a biobutanol filler, that is to say a butanol filler produced from renewable resources derived from biomass. Said butanol feed mainly comprises butanol, at a level of more than 50% by weight and preferably more than 70% by weight and also comprises water, at a content of between 0 and 50% by weight, preferably between 0% and 30% by weight. % by weight, and preferably between 15% and 25% by weight, a content of cationic impurities such as, for example, Na ⁺ , Ca ²⁺ , K ⁺ , Mn ²⁺ , Fe ²⁺ , Cu ²⁺ , Zn ² ions. ⁺ preferably less than 0.1% by weight, an anionic impurities content such as, for example chloride ions, sulphate, nitrite, nitrate, phosphate advantageously less than 0.1% by weight, a metal content such as nickel, chromium and the potassium advantageously less than 0.1% by weight, a content of other alcohols such as, for example, methanol and ethanol, advantageously less than 10% by weight, and preferably less than 5% by weight, a content of oxygenated compounds other than alcohols such as, for example, ethers, acids, ketones, aldehydes, acetals and esters advantageously less than 1% by weight and a content of nitrogen compounds and sulfur compounds such as, for example, amines, acetonitriles, nitric sulphates, carbon sulphide advantageously less than 0.5 % weight, the percentages weight being expressed relative to the total mass of said butanol feed.

The process according to the invention advantageously comprises a purification step carried out prior to the dehydration step a) so as to eliminate the cationic and anionic impurities as well as at least a part of the oxygenated compounds in order to limit the deactivation of the dehydration catalyst placed downstream.

Said purification step is advantageously carried out by means known to those skilled in the art, such as for example the use of at least one resin, the adsorption of impurities and oxygenates on solids selected from sieves Molecules, activated carbon, alumina and zeolites and distillation to produce a purified butanol feed responding to the levels of impurities compatible with the dehydration catalyst and a product comprising organic impurities.

Said levels of impurities compatible with the dehydration catalyst are an ionic impurities content of less than 100 ppm, a metal content of less than 50 ppm, an oxygenated impurities content of less than 150 ppm and a nitrogen and sulfur compounds content of less than 30 ppm. A pretreatment step may also advantageously be carried out by hydrogenation of the oxygenated unsaturated compounds in the presence of a nickel-based catalyst, said pretreatment step being carried out before or after the purification step and preferably after.

Step a) Dehydration

According to the invention, the butanol feedstock, optionally purified, undergoes a step a) of isomerizing dehydration so as to produce a predominantly butylenic effluent, said step operating in the presence of a dehydration catalyst known to those skilled in the art, in particular of an amorphous acid catalyst or a zeolitic acid catalyst in at least one reactor.

The term "predominantly butylenic effluent" means an effluent comprising water and at least 95%, preferably at least 97% and preferably at least 98% by weight of butenes relative to the total mass of the carbon compounds present in said effluent. produced by said step a). In addition to the majority presence of butenes, said carbonaceous effluent may also comprise other hydrocarbon, hydroxycarbon or oxycarbon compounds in a very minor proportion. In particular, said carbonaceous effluent advantageously comprises less than 5%, preferably less than 3% and preferably less than 2% by weight of compounds having a number of carbon atoms greater than or equal to 5 and of oxygenated compounds such as by examples C0 ₂ , CO, diethylether and acetaldehyde, the percentages being expressed in weight percentages relative to the total mass of the carbon compounds present in said effluent produced in said step a).

Said step a) of isomerizing dehydration makes it possible to convert isobutanol into a mixture of butenes.

In the case where the catalyst used in the dehydration step a) is a zeolite catalyst, said catalyst comprises at least one zeolite chosen from zeolites having at least pore openings containing 10 or 12 oxygen atoms (10MR or 12 MR). It is known to define the pore size of zeolites by the number of oxygen atoms forming the annular section of the zeolite channels, called "member ring" or MR in English. Preferably, said zeolite catalyst comprises at least one zeolite having a structural type selected from structural types MFI, FAU, MOR, FER, and BEA.

The zeolite used in the catalyst used in step a) of the process according to the invention can advantageously be modified by dealumination or desilication according to any method of dealumination or desilication known to those skilled in the art. In the case where the catalyst used in the dehydration step a) is an amorphous acid catalyst, said catalyst comprises at least one porous refractory oxide chosen from alumina, alumina activated by a mineral acid deposit and silica. alumina.

Said amorphous or zeolitic dehydration catalyst used in step a) of the process according to the invention may also advantageously comprise at least one matrix of oxide type also called binder. A matrix according to the invention is understood to mean an amorphous or poorly crystallized matrix.

Said matrix is advantageously chosen from the elements of the group formed by clays (such as, for example, natural clays such as kaolin and bentonite), magnesia, aluminas, silicas, silica-aluminas, aluminates, titanium oxide, boron oxide, zirconia, aluminum phosphates, titanium phosphates, zirconium phosphates, and coal. Preferably, said matrix is chosen from the elements of the group formed by aluminas, silicas and clays.

In a preferred embodiment, the binder has a macroporous texture as described in US Pat. No. 7,880,048.

The dehydration catalyst used in step a) of the process according to the invention is advantageously shaped in the form of grains of different shapes and sizes. It is advantageously used in the form of cylindrical or multi-lobed extrusions such as bilobed, trilobed, straight-lobed or twisted, but can optionally be manufactured and used in the form of crushed powder, tablets, rings, beads, wheels, spheres. Preferably, said catalyst is in the form of extrudates or beads.

The isomerizing dehydration step a) of the process according to the invention advantageously operates at an absolute pressure at the reactor inlet of between 0.5 and 1.2 MPa, preferably between 0.6 and 1.1 MPa and at a temperature comprised between between 350 and 450 ° C. The absolute pressure at the reactor inlet is chosen so that the said feedstock of said step a) is in the gas phase at the reactor inlet.

The high operating pressure, that is to say greater than 0.5 MPa, preferably greater than 0.6 MPa, advantageously allows the products to separate water by liquid / liquid decantation. The hourly volume velocity is between 2 and 7 h ^-1 .

By hourly volume velocity is meant the flow rate of the feedstock at the reactor inlet in m ³ / h at 15 ° C., 1 atm divided by the volume of catalyst in m ³ contained in the reactor. Said step a) is very endothermic. It is therefore advantageously operated in at least two separate reactors, the effluent of a reactor being heated before being sent as the next reactor charge. The temperature at the inlet of each reactor is between 350 and 450 ° C, preferably between 350 and 400 Ό, more preferably between 350 and 375 ° C.

The conversion of the butanol feedstock in step a) is advantageously greater than 95%, preferably 99% and preferably greater than 99.8%. By conversion of the butanol feedstock is meant the ratio of the difference between the mass flow rate of the butanol feedstock at the inlet of step a) and the mass flow rate of the butanol feedstock at the outlet of step a) on the mass flow rate. butanol feed at the inlet of step a).

Step b) Separation of water

According to the invention, the predominantly butylenic effluent from step a) undergoes at least one step b) of separating the water present in said effluent. The water has a detrimental effect on the catalysts of the subsequent steps of the process according to the invention. The effluents of step b) are an aqueous liquid effluent, an organic liquid effluent and optionally a gaseous effluent comprising carbon monoxide, carbon dioxide, hydrogen, methane and acetone, taken alone or in mixture. Said step b) is preferably a decantation step in which an aqueous phase is separated from an organic phase.

Said step b) advantageously operates at a pressure of between 0.5 and 1.2 MPa, preferably between 0.6 and 1.1 MPa and at a temperature between 35 and 60 ° C.

The high operating pressure of the dehydration step a) according to the invention makes it possible to condense the water in step b) at an elevated temperature. When the dehydration step is carried out at a lower pressure, it is either necessary to use a cold cycle to condense the water, or it is necessary to compress the effluent at the outlet of the dehydration step.

Step c) Purification

According to the invention, the organic liquid effluent from step b) undergoes a purification step c) so as to produce a purified organic effluent.

Said step c) may advantageously be carried out by any method known to those skilled in the art, for example by treatment in an MDEA (methyldiethylamine) or other amine absorption column followed by a molecular sieve treatment, said sieves being advantageously of the 13X, 3A, 4A and 5A type, taken alone or as a mixture. Said step c) can also advantageously be carried out by treatment in a washing column with sodium hydroxide. Said step c) can also advantageously be implemented by treatment of said organic effluent on sieves molecular, said sieves being advantageously of the 13X, 3A, 4A and 5A type, taken alone or as a mixture. Said step c) may also advantageously comprise a treatment on basic resin so as to remove the alcohols present in said organic liquid effluent. Dryers may advantageously be implemented so as to achieve a water content compatible with the oligomerization catalysts used downstream in steps d) and e) of oligomerization.

The water content of said purified organic effluent is advantageously between 0 and 1000 ppm, preferably between 0 and 500 ppm and preferably between 0 and 200 ppm. The content of sulfur-containing constituents, for example H ₂ S or COS, of said purified organic effluent is advantageously less than 100 ppm, preferably less than 50 ppm. The content of nitrogen components, for example ammonia, of said purified organic effluent is less than 1 ppm. The content of inorganic constituents of said purified organic effluent is less than 1 ppm, preferably less than the detection limits.

Step d) selective oligomerization of isobutenes called "Selectopol"

According to the invention, at least a portion of the purified organic effluent from purification step c), all of the effluent from oligomerization step g) and at least a portion of the light product. from fractionation step f) are mixed so as to form a charge of first oligomerization, said feedstock undergoing a first step d) selective oligomerization of isobutenes in at least one reactor in the presence of an amorphous catalyst so as to produce an effluent of first oligomerization.

By at least a portion of the purified organic effluent from step c), is meant at least 50% by weight, preferably at least 90% by weight and preferably all of the effluent from said step c) , the weight percentages being expressed relative to the total mass of said effluent.

The conversion of isobutene in step d) is advantageously greater than 80%, preferably greater than 85% and preferably greater than 90%. Isobutene conversion is understood to mean the ratio of the difference between the mass flow rate of isobutene at the inlet of step d) and the mass flow rate of isobutene at the outlet of step d) on the mass flow rate of isobutene at the input of step d).

The conversion of the n-butenes in step d) is advantageously less than 15%, preferably less than 10% and preferably less than 5%. The conversion of n-butenes is defined analogously to the conversion of isobutene.

Said effluent of first oligomerization comprises at least 50% by weight of olefins having a number of carbon atoms greater than or equal to 8, the weight percentage being expressed relative to the total mass of olefins contained in said effluent. In particular, said effluent comprises at least at least 60% by weight, preferably at least 65% and preferably at least 75% by weight, of olefins having a number of carbon atoms greater than or equal to 8, the weight percentage being expressed relative to the total mass olefins contained in said effluent. Of the olefins having a number of carbon atoms greater than or equal to 8, the C ₈ -C ₁₂ olefins are in the majority with respect to the olefins having at least 13 carbon atoms, that is to say that the mass ratio C ₈ -C ₁₂ olefins on olefins having at least 13 carbon atoms is greater than 1.

In addition to the majority presence of olefins having a number of carbon atoms greater than or equal to 8, the olefinic effluent produced during the first oligomerization step d) also advantageously comprises less than 50%, preferably less than 40% and preferably, less than 35% and very preferably less than 25% by weight of olefinic compounds whose number of carbon atoms is less than or equal to 4, the weight percentages being expressed relative to the total mass of the olefins contained in the effluent produced.

The recycling of all the effluent from the oligomerization step g) and at least a portion of the light product from the fractionation step f) makes it possible to improve the conversion efficiency of the olefins having less of 8 carbon atoms by better control of the exothermicity.

Olefins comprising at least 16 carbon atoms react very little in step d) and in step e). Their recycling upstream of step d) makes it possible to minimize the rise in temperature, either in the reaction section of step d) or in the reaction section of step e). This recycling therefore allows better temperature control in said reaction sections.

The catalyst used in the first oligomerization step d) comprises at least one group VI II element, preferably chosen from nickel, cobalt, iron, platinum and palladium, and preferably, said element is nickel and at least one porous oxide refractory support preferably chosen from alumina, silica, silica-aluminas, siliceous aluminas, zirconia, titanium oxide, magnesia, clays taken alone or as a mixture, and preferably, said support is an alumina or a silica alumina, preferentially a silica-alumina. In a preferred arrangement, the catalyst used in the first oligomerization step d) is a silica-alumina catalyst as described in US Pat. No. 7,572,946.

Said catalyst used in said step d) of the process according to the invention also advantageously comprises at least one matrix of oxide type also called binder. A matrix according to the invention is understood to mean an amorphous or poorly crystallized matrix.

Said matrix is advantageously chosen from the elements of the group formed by clays (such as for example from natural clays such as kaolin or bentonite), magnesia, aluminas, silicas, silica-aluminas, aluminates, titanium oxide, boron oxide, zirconia, aluminum phosphates, titanium phosphates, zirconium phosphates, and charcoal. Preferably, said matrix is chosen from the elements of the group formed by alumina, clays and silicas, more preferably said matrix is chosen from aluminas, and even more preferably said matrix is gamma-alumina.

The catalyst used in said step d) of the process according to the invention is advantageously shaped in the form of grains of different shapes and sizes. It is advantageously used in the form of cylindrical or multi-lobed extrusions such as bilobed, trilobed, straight-lobed or twisted, but can optionally be manufactured and used in the form of crushed powder, tablets, rings, beads, wheels, spheres. Preferably, said catalyst is in the form of extrudates of size between 1 and 10 mm.

Said first oligomerization step d) of the process according to the invention is advantageously carried out at a temperature of between 40 and 110 ° C., preferably between 60 and 90 ° C., at an absolute pressure of between 0.5 and 10 MPa. preferably between 1 and 8 MPa and preferably between 2 and 6 MPa, chosen so as to maintain products and reagents in liquid form, and at an hourly space velocity of between 0.1 and 10 h ^-1 and preferably between 0 and 10 h ^-1. , 4 and 5 hr ^-1 .

Said first d) oligomerization step of the process according to the invention is preferably carried out in a fixed bed. Preferably, said step is performed in two fixed bed reactors in series.

Step e) oligomerization of olefins called "Polynaphtha"

According to the invention, the first oligomerization effluent from step d) undergoes a second oligomerization step e) in the presence of an amorphous catalyst so as to produce a second oligomerization effluent.

The second oligomerization step e) allows the production of an effluent enriched in olefins having a carbon number greater than or equal to 9.

Said effluent of second oligomerization is an olefinic effluent advantageously comprising less than 50% by weight and preferably less than 45% by weight of olefins having a number of carbon atoms of between 4 and 8, the weight percentages being expressed relative to the total mass of the olefinic effluent C ₄ -C ₈ entering said second oligomerization step e).

The conversion of isobutene in step e) is advantageously greater than 80%, preferably greater than 85%, preferably greater than 90% and very preferably greater than 95%. The conversion of n-butene in step e) is advantageously greater than 60%, preferably greater than 70%, and preferably greater than 80%.

The conversion of olefins having a number of carbon atoms of between 8 and 11 in step e) is advantageously greater than 40%, preferably greater than 50%, and preferably greater than 60%.

The conversion of olefins having a number of carbon atoms of at least 12 in step e) is advantageously less than 5%, preferably less than 3%, and preferably less than 1%.

The temperature rise in the reactors due to the exothermicity of the reactions in steps d) and e) will be controlled by varying the recycling rate of the light product from step f) as well as the fraction of the intermediate product from from step f) sent to step g).

The n-butenes and olefins having a carbon number of at least 8 being unreactive in the first step d) oligomerization, they play the role of thermal fly in said step d).

Olefins having a carbon number of at least 12 being unreactive in the second step e) oligomerization, they play the role of thermal fly in said step e).

The catalyst used in step e) of the process according to the invention has the same characteristics as the catalyst used in step d) of the process according to the invention. Preferably, it is identical to the catalyst used in step d) of the process according to the invention.

Said second oligomerization step e) of the process according to the invention is advantageously carried out at a temperature of between 100 and 200 ° C., preferably between 10 and 160 ° C., at an absolute pressure of between 2 and 15 MPa, preferably between 2 and 15 MPa. and 8 MPa and preferably between 3 and 8 MPa and at an hourly space velocity of between 0.1 and 10 h ^-1 and preferably between 0.4 and 5 h ^-1 . In any case, the operating pressure of the process is such that all the reagents and products are in liquid form in the reaction zone.

Said oligomerization step e) is advantageously carried out in at least one fixed bed reactor, preferably at least two fixed bed reactors in series, and more preferably at least three fixed bed reactors in series. The operation of step e) at a high temperature allows the production of less connected molecules than if this step was operated at lower temperature, which leads to obtaining a product whose density is higher.

Step f) Splitting

According to the invention, the second oligomerization effluent from step e) undergoes a fractionation step f) in at least one distillation column so as to separate said effluent into at least three products respectively corresponding to a light product. comprising predominantly C ₂ to C ₄ compounds, an intermediate product comprising predominantly C ₅ to C ₉ compounds, corresponding to a gasoline cut and a middle distillate product comprising predominantly compounds having at least 10 carbon atoms whose cutting point is between 150 and 350 ° C. A bottoms product having an initial boiling point of 350 to 370 ° C is also preferably separated. These products are not restrictive.

Said light product advantageously comprises at least 80% by weight of n-butenes, preferably 85% by weight, and preferably 90% by weight, the weight percentage being expressed relative to the total amount of butenes present in said light product.

At least a portion of said light product may advantageously be recycled in the first oligomerization step d) of the process according to the invention.

By at least a portion of said light product is advantageously meant between 0 and 100% by weight of the total mass flow rate of said light product, preferably between 50 and 100% by weight, more preferably between 75 and 100% by weight.

At least a portion of said intermediate product is treated in an oligomerization step g).

By at least a part of said intermediate product is advantageously meant between 80 and 100% of the total mass flow rate of said intermediate product, preferably between 90 and 100%, more preferably between 95 and 100%.

According to the invention, at least a portion of said middle distillate product is treated in a hydrogenation step (h). By a part of said middle distillate product is meant at least 80% of the total flow rate of the middle distillate product, preferably at least 90% and preferably all of said middle distillate product.

Said step f) is advantageously carried out with two distillation columns operating in series, the first distillation column fractionating the effluent resulting from the second oligomerization step e). a light product comprising predominantly C ₂ to C ₄ compounds and a first column bottoms product, and the second column fractionating said first column bottoms product to an intermediate product comprising predominantly C ₅ to C ₉ compounds and a distillate product medium composed mainly of compounds having at least 10 carbon atoms.

Said light product advantageously comprises at least 90% by weight of olefins whose number of carbon atoms is less than or equal to 4, preferably at least 95% by weight, more preferably at least 98% by weight and preferably at least 99%, 9% by weight, this percentage being defined relative to the total weight of said light product.

Said intermediate product advantageously comprises at least 90% by weight of olefins whose number of carbon atoms is between 5 and 9, preferably at least 96% by weight and more preferably at least 99.8% by weight, this percentage being defined with respect to total weight of said intermediate product.

Said middle distillate product advantageously comprises at least 90% by weight of olefins whose number of carbon atoms is at least equal to 10, preferably at least 95% by weight, more preferably at least 98% by weight and, preferably, at least 98% by weight. 99.5% by weight, this percentage being defined relative to the total weight of said middle distillate product.

Stage g) oligomerization of C ₈ olefins called "C ₈ Polynaphtha"

According to the invention, at least a portion of said intermediate product produced by step f) of the process according to the invention is treated in a third step g) oligomerization.

The conversion of olefins having 8 carbon atoms in step g) is advantageously between 20 and 40%. The selectivity of the oligomerization reaction of olefins having 8 carbon atoms to olefins having 16 carbon atoms is greater than 95%, preferably greater than 98%, preferably greater than 99%.

The catalyst used in step g) of the process according to the invention has the same characteristics as the catalyst used in step d) of the process according to the invention. Preferably, it is identical to the catalyst used in step d) of the process according to the invention.

Said oligomerization step g) of the process according to the invention is advantageously carried out at a temperature of between 100 and 200 ° C., preferably between 110 ° and 160 ° C., at an absolute pressure of between 2 and 15 MPa, preferably between 2 and 8. MPa and preferably between 3 and 8 MPa and at an hourly space velocity of between 0.1 and 5 h ^-1 and preferably between 0.4 and 2 h ^-1 . In In any case, the operating pressure of the process is such that all the reagents and products are in liquid form in the reaction zone.

Said oligomerization step g) is advantageously carried out in at least one fixed bed reactor, preferably at least two fixed bed reactors in series, and more preferably at least three fixed bed reactors in series.

Step h) Hydrogenation

According to step h) of the process according to the invention, at least a portion of said middle distillate product from step f) undergoes a step of hydrogenation of olefins into paraffins, to make them incorporable to the fuel pool.

Preferably, at least a portion and preferably all of said middle distillate product from step f) is contacted with a hydrogen-rich gas in the presence of a catalyst comprising at least one Group VII metal, advantageously chosen from palladium and nickel taken alone or as a mixture, and a support advantageously chosen from alumina, silica or silica-alumina.

The catalyst used in said hydrogenation step h) comprises a palladium content of between 0.1 and 5% by weight and / or a nickel oxide content advantageously between 15 and 40% by weight relative to the total mass. catalyst. This nickel can be promoted with molybdenum or can also be partially sulphurized. As the metal content of the catalyst increases, the operating temperature of the reaction section must be adjusted downward to compensate for the increase in catalytic activity.

The hydrogenation step h) advantageously operates at a temperature of between 100 and 350 ° C. at the reactor inlet and at a pressure of between 2 and 4 MPa and at an hourly space velocity of between 1 and 5 h ^-1 . The molar ratio of hydrogen to H ₂ / HC hydrocarbons is between 10 and 450.

The unreacted hydrogen is separated at the outlet of the reactor in order to be recycled as input. A fraction of the total flow of the reaction effluent freed from hydrogen is also recycled at the inlet so as to act as a thermal flywheel for the reaction. The flow rate of said recycled fraction represents between 2 and 5 times the mass flow rate of said portion of said middle distillate product from step f) at the inlet of step h), that is to say before the recycling of the hydrogen and said fraction.

The reactors used in step h) are multi-bed reactors with a recycling of a portion of the reaction effluent to dilute the charge and control the exotherm. Hydrogenation concerns almost exclusively olefins, the reaction is therefore very exothermic (between 100 and 300Ό of exotherm)

The performance of the hydrogenation is validated by a measurement of the smoke point and current gums which will advantageously be greater than 25 mm for the smoke point and less than 7 mg / 100 ml for the current gum content. This generally results in a measurement of the number of bromine according to ASTM D271 0 which is advantageously at most 10 mg Br / 100g when these limits for the smoke point and the current gums are met.

The effluent from the optional hydrogenation stage mainly contains valuable hydrocarbons that can be incorporated into the kerosene and / or diesel fuel pool and preferably kerosene.

The yield of hydrocarbons in which the number of carbon atoms is at least equal to 10 of said hydrogenation step h) is greater than 90%, preferably greater than 95%. The olefin content in the effluent of said step h) is between 0 and 5% of the total weight of said effluent, preferably between 0 and 2% by weight.

An optional separation step according to the hydrogenation step h) is advantageously used to allow the fractionation into a kerosene cut and / or a gas oil cut and / or a cut having a boiling point greater than 360 ° C. and / or light cuts.

EXAMPLES

Example 1 (compliant)

This example illustrates the invention.

Description of Butanol Filler

The butanol feed used in the example was treated by a succession of distillation steps and molecular sieve passes. Following these treatments, the purified feed has the composition indicated in Table 7, column "load step a)". The purified feedstock has an ionic impurity content of less than 100 ppm, a metal content of less than 50 ppm, an oxygenated impurity content of less than 150 ppm and a total content of nitrogen and sulfur compounds of less than 30 ppm.

Step a): Dehydration of the purified feed

The purified feed undergoes a step a) of isomerizing dehydration. Said step a) is carried out at a temperature of 400 ° C., at a pressure of 0.85 MPa at the reactor inlet, and at a speed of hourly volume of 5 h ^-1 in the presence of a silica-alumina catalyst so as to maximize the production of butenes.

The effluent of step a), mainly butylenic, has the composition indicated in Table 7, column "effluent step a)". Isobutanol is converted to 99.8% to a mixture of butenes close to thermodynamic equilibrium and water. The reaction also produces a number of co-products: oxygenates and light olefins.

Step b): separation of water from the butylenic effluent

The predominantly butylenic effluent from step a) is then directed to a settling tank to effect the separation of the water. The separation is carried out at 55 ° C. and 0.75 MPa so as to promote the liquid-liquid demixing and thus the separation of the hydrocarbons from the water. The residual water content in the organic liquid effluent is then 1800 ppm.

Table 1 - Mass Distribution of Hydrocarbon Compounds in the Effluent of Step b)

The fraction of oxygenates, nitrogen and water in the organic liquid effluent from step b) is 0.3%. .

Step c): purification of the organic liquid effluent

The organic liquid effluent from step b) then goes through a purification step in order to remove most of the remaining water and the compounds that may affect the downstream catalytic beds. This purification step allows the remaining sulfur, nitrogen and oxygen elements to be partially removed on a 13x molecular sieve, and then to partially remove the remaining water on a 3A molecular sieve.

The composition of the purified organic effluent leaving step c) is indicated. TABLE 2

Table 2 Step d): step of selective oligomerization of isobutene

The purified organic effluent from step c) is mixed with 76.5% of the total flow of the top effluent from fractionation step f) and all the effluent from step g). oligomerization of C ₈ olefins. This mixture has the composition shown in Table 7, column "charge stage d)". It is treated in the first oligomerization step d), which operates in the presence of the catalyst IP81 1 marketed by Axens. Catalyst IP81 1 is an amorphous silica-alumina catalyst. The operating conditions of stage d) are a temperature of 60.degree. C., a pressure of 3 MPa at the reactor inlet and a hourly space velocity in the reactors of 2 h.sup.- ¹ so as to favor, for the most part, the oligomerization of the reactor. isobutene while limiting the reaction of n-butenes.

The effluent of the oligomerization step d) has the composition indicated in Table 7, column "charge step e)".

After step d), from 1 kg of butanol introduced in step a), 295 g of olefins having a number of carbon atoms greater than or equal to 5 are produced.

Step e): Oligomerization step of olefins

The effluent from step d) of selective oligomerization then treated in step e). The latter is operated in the presence of the catalyst IP81 1 marketed by Axens. The operating conditions of step e) are a temperature of 1 1 0Ό, a pressure of 6 MPa at the reactor inlet and a speed rate per hour in the reactors of 2 h ^-1 .

The effluent of the oligomerization step e) has the composition indicated in Table 7, column "charge step f)".

Step f): Fractionation step of the oligomerization effluents

The effluent of step e) is then subjected to fractionation step f) so as to separate a light product comprising the compounds C ₂ to C ₄ , an intermediate product comprising the compounds C ₅ to C ₉ and a distillate product. medium, which will be partly the kerosene cut, composed of C ₁₀ ⁺ . The breakdown of the separated products is presented Table 3.

Table 3 - Distribution of Products at the End of Step f)

Starting from 1 kg / h of butanol at the inlet of stage a), 63 g / h of light product, 618 g / h of intermediate product and 702 g / h of medium distillate product are obtained at the outlet of the product. step f). A large part of first two products is recycled to obtain these results: 76.5% of the light product is recycled at the input of step d) and 96.5% of the intermediate product is directed to step g) in order to oligomerize the olefins C ₅ -C ₈ before being redirected to step d).

The light product contains 500 ppm mass of carbon compounds other than C ^ C ^ The intermediate product contains 2000 ppm mass of carbon compounds other than C ₅ -C ₉ . The middle distillate product contains 0.5% mass of carbon compounds

Step g): oligomerization of C8 olefins

96.5% of the intermediate product from step f) is directed to step g) of oligomerization of C ₈ olefins. This step is carried out in the presence of the catalyst IP81 1 marketed by Axens. The operating conditions of step g) are a temperature of 110 ° C., a pressure of 6 MPa at the reactor inlet and a speed rate per hour in the reactors of 2 h ^-1 .

The composition of the effluent leaving step g) is presented in Table 4. It is then recycled at the inlet of step d).

Table 4

Step h): Hydrogenation of olefins

The middle distillate product resulting from the fractionation stage f) is directed to the hydrogenation stage olefins which operates with an LD746 catalyst marketed by Axens and in which the majority of the olefins will be hydrogenated while minimizing the production of light hydrocarbon molecules. This step operates at a temperature of 160 Ό and a pressure of 2.5 MPa at the reactor inlet. The hourly volume velocity in the reaction section is 3h ^"1 and the H ₂ / HC ratio is 100.

Step h) produces a mixture of alkanes having predominantly between 9 and 16 carbon atoms, the distribution of which is shown in Table 5.

The effluent from the hydrogenation step h) is then fractionated and the corresponding fractions are then mixed with the non-recycled fractions of light and intermediate products from step f).

The distribution of the products resulting from the process operated according to Example 1 is presented in Table 6.

step h), from 1 kg / h of butanol entering step a), 561 g / h of bio-kerosene are

Table 7 - Main Streams of Example 1

Example 2 (non-compliant) - Production process without oligomerization of the recycled intermediate product

In this example, the intermediate product from step f) is recycled to the input of step d) without oligomerization in a step g). The charge, as well as steps a) and b) and c) are identical to Example 1. Step d): oligomerization step of isobutene

The predominantly butylenic effluent from step c) is then mixed with 77% of the light product and 87.5% of the intermediate product from the fractionation step f) and then sent to the first oligomerization step d), which is operated under the same operating conditions and with the same catalyst as in Example 1.

The composition of the effluent at the outlet of the reaction is presented in Table 8.

Table 8

At the end of this step, from 1 kg / h of butanol entering in step a), 295 g / h of olefins having more than 5 carbon atoms are produced.

The effluent of step d) is then directed to step e) of oligomerization of olefins. Step e): Oligomerization step of olefins

Step e) operates under the same conditions and with the same catalyst as Example 1. The composition of the second oligomerization effluent is indicated in Table 9.

Table 9 Step f): Fractionation step of the oligomerization effluents

The effluent from step e) is then fractionated into 3 products in step f), in the same manner as in Example 1. The breakdown of the separated products is presented in Table 10.

Table 10

Starting from 1 kg / h of butanol at the inlet of stage a), 61 g / h of light product, 733 g / h of intermediate product and 633 g / h of medium distillate product are obtained. 77% mass of the light product and 87.5% mass of the intermediate product are recycled in input in step d).

The light product contains 500 ppm mass of carbon compounds other than C ^ -C ^. The intermediate product contains 2000 ppm mass of carbon compounds other than C ₅ -C ₉ . The middle distillate product contains 0.5% mass of carbon compounds

Step h): Hydrogenation of olefins

The middle distillate product from the fractionation step f) is directed to the hydrogenation step β of the olefins. Step h) operates under the same conditions and with the same catalyst as Example 1. The effluent of the hydrogenation has the composition shown in Table 1 1.

The distribution of the products resulting from the process operated according to Example 2 is presented Table 12.

At the exit of step h), from 1 kg / h of butanol entering in step a), 490 g of bio-kerosene are obtained.

Example 3 (non-compliant) - Production process without intermediate product recycling

In this example, neither the intermediate product nor the light product from the fractionation step f) is recycled to the input of the oligomerization step d). In contrast, the intermediate product undergoes an oligomerization step g) to improve the reaction yield.

The charge, as well as steps a), b) and c) are identical to Example 1.

Step d): oligomerization step of isobutene

The predominantly butylenic effluent from step c) is then sent to the first oligomerization step d), carried out under the same operating conditions and with the same catalyst as in Example 1. The composition of the effluent at the outlet of the reaction is shown in Table 13.

Table 13

At the end of this step, from 1 kg / h of butanol entering in step a), 294 g / h of olefins having more than 5 carbon atoms are produced.

Step e) operates under the same conditions and with the same catalyst as Example 1. The composition of the olefinic effluent of step e) is presented in Table 14. The mass distribution of the

hydrocarbon compounds% wt

butane 0.4%

C2 = 0.7%

iC4 = 5.9%

Table 14

Step f): Fractionation step of the oligomerization effluents

The effluent from step e) is then fractionated into 3 products in step f), in the same manner as in Example 1. The breakdown of the separated products is presented in Table 15.

Table 15

From 1 kg / h of butanol at the inlet of step a), 52 g / h of light product, 295 g / h of intermediate product and 400 g / h of medium distillate product are obtained. No product is recycled in this example, contrary to what is done in Examples 1 and 2. All of the intermediate product is directed to step g) oligomerization of C ₈ olefins to increase the final yield . The light product is integrated as is in the products. The middle distillate product is directed to the hydrogenation step h).

Step g): oligomerization of C ₈ olefins

The intermediate product from fractionation step f) is directed to step g). Step g) operates under the same conditions and with the same catalyst as in Example 1. The composition of the effluent leaving step g) is indicated. Table 16.

Table 16 It is then directed to the hydrogenation step olefins to reduce its olefin content together with the average distillate product from the fractionation step f).

Step h): Hydrogenation of olefins

The middle distillate product from the fractionation step f) and the effluent from step g) are directed to the hydrogenation step β of the olefins. Step h) operates under the same conditions and with the same catalyst as Example 1. The effluent of the hydrogenation has the composition shown in Table 17.

The effluent from the hydrogenation step h) is then fractionated and the corresponding fractions are then mixed with the light and intermediate product fractions from step f).

The output composition of the hydrocarbon effluent of the process is indicated in Table 1.

At the end of step h), from 1 kg / h of butanol entering in step a), 402 g of bio-kerosene are obtained.

Example 4 (non-compliant) - Production method without recycling of the intermediate product or oligomerization of said intermediate product

This example is similar to Example 3. However, the intermediate product does not undergo step g) of oligomerization before being incorporated into the fuel pool.

The charge, as well as steps a) to f) are identical to Example 3. The middle distillate product from step f) is hydrogenated in step h) of hydrogenation of the olefins. The light and intermediate products from step f) are directly integrated into the products of the process.

Step h): Hydrogenation of olefins

The middle distillate product from the fractionation step f) is directed to the hydrogenation step β of the olefins. Step h) operates under the same conditions and with the same catalyst as in Example 1.

The effluent of the hydrogenation has the following composition shown in Table 19.

Table 19

The output composition of the hydrocarbon effluent of the process is indicated in Table 20.

At the end of step h), from 1 kg / h of butanol entering in step a), 319 g of bio-kerosene are obtained.

Example 5 (non-compliant) - Production process without selective oligomerization of isobutenes, with oligomerization of the intermediate product and recycling

In this example, butene oligomerization is carried out in one step and not in two steps d) and e) as taught in the invention. The i-butenes, which are more reactive, then react at the same time as the other olefins comprising from 4 to 8 carbon atoms.

The charge, as well as steps a), b) and c) are identical to Example 1. Oligomerization stage of olefins

The predominantly butylenic effluent from step c) is mixed with 35.7% of the light product from step f) and all of the effluent from step g) and then sent to step d). oligomerization, carried out in the presence of the catalyst IP81 1 marketed by Axens, at a temperature of 60 ° C., a pressure of 3 MPa at the reactor inlet and an hourly space velocity of 2 h ^-1 . The composition of the olefinic effluent of the oligomerization step is indicated in Table 21.

Table 21

At the end of this step, from 1 kg / h of butanol entering in step a), 937 g / h of olefins having more than 5 carbon atoms are produced.

Step f): Fractionation step of the oligomerization effluents

The effluent from the oligomerization step is then fractionated into 3 products in step f), in the same manner as in Example 1. The breakdown of the separated products is presented Table 22.

Table 22

From 1 kg / h of butanol at the inlet of step a), 93 g / h of light product, 576 g / h of intermediate product and 695 g / h of medium distillate product are obtained.

35.7% of the light product is recycled upstream of the oligomerization step and 100% of the intermediate product is directed to step g) in order to oligomerize the C ₅ -C ₈ olefins before being recycled to the oligomerization step. The light product contains 500 ppm mass of carbon compounds other than C ₁ -C ₄ - The intermediate contains 2000 ppm mass of carbon compounds other than C ₅ -C ₉ . The middle distillate product contains 0.5% mass of carbon compounds

Step g): oligomerization of olefins C8 =

The intermediate product from fractionation step f) is directed to step g) of oligomerization of C ₈ olefins. Step g) operates under the same conditions and with the same catalyst as in Example 1. The composition of the effluent leaving step g) is indicated in Table 23.

Table 23

The effluent from step g) is then recycled to the oligomerization step. Step h): Hydrogenation of olefins

The middle distillate product from the fractionation step f) is directed to the hydrogenation step β of the olefins. Step h) operates under the same conditions and with the same catalyst as in Example 1. The effluent of the hydrogenation has the composition shown in Table 24.

The effluent from the hydrogenation step h) is then fractionated and the corresponding fractions are then mixed with the non-recycled light and intermediate product fractions from step f).

The output composition of the hydrocarbon effluent of the process is indicated in Table 25.

At the exit of step h), from 1 kg / h of butanol entering in step a), 547 g of bio-kerosene are obtained.

Synthesis of the results of Examples 1 to 5

Table 26 compares some key results of Examples 1 to 5. The bio-kerosene yield corresponds to the amount of bio-kerosene produced per 1 kg of butanol in the process feed.

^* The yields include a 5% loss in the hydrogenation of the kerosene cut (step h).

Table 26

Example 1 shows that the exotherm in both steps d) and e) is significantly smaller and the kerosene yield higher than in one case without recycling or oligomerization of the intermediate product (Examples 3 and 4). Example 2 shows that the addition of step g) improves the kerosene yield.

Example 5, with a single oligomerization step, but comprising the recycling of the light and intermediate products and the oligomerization thereof, has a yield equivalent to the example 1 compliant. However, the percentage of converted butenes is lower and the control of the reaction and the products obtained are different. Indeed, the two-step separation of the oligomerization (steps d and e) makes it possible to carry out only the reactions desired in each step and thus to control both the quality of the products obtained and the exotherms in each step. The control of the exotherms in each step also makes it possible to increase the life of the catalyst (thermal stress). This impact also concerns safety (less impact of a jolt on the change in the composition of the load). Flexibility is therefore greater.

Example 6 - Aromatic Additions and Effect on Density

The kerosene product obtained in Example 1 according to a density of about 780 kg / m ³ . Mesitylene (1,3,5-trimethylbenzene) with a density of 865 kg / m ³ is added to this product.

The addition of Mesitylene in our kerosene up to an 8% vol. Raises the density of the mixture to 787kg / m3. The addition of Mesitylene in our kerosene up to a content of 25% vol. raises the density to 801 kg / m3.

It remains therefore within the density limits set by ASTM D7566-1 1 a, even by adding aromatics up to the limits provided by said standard.

Claims

A process for producing medium distillate hydrocarbon bases from a butanol feed, said process comprising at least: a) a step of isomerizing dehydration of said butanol feedstock in the presence of an amorphous or zeolitic acid catalyst in at least one reactor, operating at an absolute pressure at the reactor inlet of between 0.5 and 1.2 MPa and at a reactor inlet temperature of between 350 and 450 ° C. so as to produce a predominantly butylenic effluent,

b) a step of separating the water present in said butylene effluent operating at a pressure of between 0.5 and 1.2 MPa and at a temperature between 35 and 60 ° C,

c) a step of purifying the organic liquid effluent from step b) so as to produce a purified organic effluent,

d) a first oligomerization stage of a feedstock comprising at least a part of the purified organic effluent resulting from stage c), the totality of the effluent resulting from stage g) and at least a part of the light product resulting from step f) in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 0.5 and 10 MPa, at a temperature of between 40 and 110 Ό and at a volume rate between 0.1 and 10 h ^-1 , so as to produce an effluent of first oligomerization comprising at least 50% by weight of olefins having a number of carbon atoms greater than or equal to 8, the weight percentage being expressed relative to the total mass of olefins contained in said effluent, e) a second oligomerization step of said first oligomerization effluent, in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 2 and 15 MPa at a te mpérature between 100 and 200Ό and at a hourly volume velocity between 0.1 and 10 h ^"1 , so as to produce a second oligomerization effluent,

f) a step of fractioning said second oligomerization effluent into at least three products corresponding respectively to a light product comprising predominantly C ₂ to C ₄ compounds, an intermediate product comprising predominantly C ₅ to C ₉ compounds, and an average distillate product; comprising predominantly compounds having at least 10 carbon atoms,

g) a third step of oligomerizing at least a portion of said intermediate product in the presence of an amorphous catalyst in at least one reactor operating at an absolute pressure of between 2 and 15 MPa at a temperature of between 100 and 200 ° C and at a volume hourly velocity of between 0.1 and

5 hrs ^"1 ,

h) a step of hydrogenating at least a portion of said middle distillate product in the presence of a catalyst comprising at least one Group VI II metal in at least one reactor operating at an absolute pressure of between 2 and 4 MPa, at a temperature between 100 and 350 Ό and at an hourly space velocity of between 1 and 5 h ^-1 with a molar ratio of hydrogen on H ₂ / HC hydrocarbons of between 10 and 450.

2. The method of claim 1 wherein a step of purifying said butanol feed is performed prior to step a).

3. Method according to one of claims 1 to 2 wherein the catalyst used in step a) of dehydration is a zeolite catalyst comprising at least one zeolite selected from zeolites having at least pore openings containing 10 or 12 atoms of oxygen.

4. Method according to one of claims 1 to 2 wherein the catalyst used in step a) of dehydration is an amorphous acid catalyst comprising at least one porous refractory oxide selected from alumina, alumina activated by a deposit of mineral acid and silica alumina.

5. Method according to one of claims 1 to 4 wherein said step b) is a decantation step in which an aqueous phase is separated from an organic phase.

6. Method according to one of claims 1 to 5 wherein between 80 and 100% of the total mass flow of said intermediate product is treated in said g) oligomerization step.

7. Process according to one of claims 1 to 6 wherein at least 80% of the total flow of the middle distillate product is treated in said hydrogenation step h).

8. Method according to one of claims 1 to 7 wherein said step f) is implemented with two distillation columns operating in series.

9. Method according to one of claims 1 to 8 wherein a separation step following the hydrogenation step h) is implemented to allow fractionation into a kerosene cut and / or a diesel cut and / or a cut having a boiling point above 360 ° C and / or light cuts.